Catalysts for the conversion of alkoxy compounds, such as methanol and/or dimethyl ether, to gasoline grade hydrocarbon products have been known since the 1970's. Processes for the conversion of natural gas and/or coal to methanol have been known for a much longer period.
Since the advent of a catalyst for the conversion of methanol to gasoline as disclosed by various Mobil Oil Corporation patents issued beginning in the 1970's, worldwide only one commercial installation which practices the methanol to gasoline ("MTG") process with a Mobil-type MTG catalyst has come into being. That installation is in New Zealand.
Conceived in the late 1970's when crude oil prices were in the $30-$40/barrel range--with projections for even further price increases--the New Zealand installation then appeared to be economically feasible. As then envisioned, and as now existing, the New Zealand facility is basically a natural gas to methanol plant (as the front end) coupled to a methanol to gasoline plant operating with an MTG catalyst in a plurality of fixed catalytic bed reactors arranged in parallel, each for single pass gas flow. The arrangement comprises five MTG reaction vessels, four of which operate at any given time and one is off stream for MTG catalyst regeneration or replacement.
Between the conception of the New Zealand facility with its attendant capital and operational cost commitments and the time that the facility was completed and ready for on-stream operations (at a capital cost obligation of about $1.2 billion) the worldwide price of crude oil drastically declined to its current day levels of about $20/barrel. At today's cost of crude oil, gasoline produced from methanol by a process as followed in the New Zealand facility is competitively uneconomical in comparison to gasoline refined from crude oil. In the first place the New Zealand approach is too costly in the conversion of natural gas into methanol. The present inventor, together with Lowell D. Fraley, has described a natural gas to methanol conversion process with much lower expected capital costs in U.S. Pat. Nos. 5,177,114 and 5,245,110. Yet, even with this expected lower cost for methanol the New Zealand process of converting methanol to gasoline is still too high in capital cost.
A major reason that the methanol to gasoline part of the New Zealand facility is uneconomical is because of the heat exchange duties--both in capital and operational cost--required for proper operation of a Mobil-type methanol to gasoline catalyst (MTG catalyst). For proper operation of an MTG catalyst in terms of its activity and aging, it must be exposed to a particularly controlled environment of temperature and water concentration while it is in contact with the methanol (or other oxygenated compounds) that it converts to gasoline grade hydrocarbons. If exposed to too low of a temperature, an MTG catalyst will not be active or will be of such low catalytic activity to be of no interest. Exposed to too high of a temperature, the MTG catalyst will so prematurely "age" and/or be destroyed so as to be economically impractical for use.
This problem of the temperature sensitivity of an MTG catalyst is compounded by the problem of its need for the presence of water in order for it to be catalytically active. Too little water in its presence and it is inactive or of such low activity to be of little practical interest. Depending upon its temperature exposure, too high a concentration of water and the MTG catalyst ages so prematurely and/or is destroyed so as to render its economical use unfeasible.
These temperature/water sensitivity problems attendant to use a MTG catalyst are further complicated by the fact that the reaction of methanol (or other oxygenated compounds) to gasoline which it catalyzes is fast and highly exothermic, thus causing a high degree of heat release (i.e., heat of reaction) at the site of the desired reaction. If not moderated by some means, a methanol and/or dimethyl ether feedstock fed into contact with an MTG catalyst at 710.degree. F. would react in contact with the catalyst to form a product gas having a temperature greater than 2,000.degree. F. Such a temperature rise within an MTG reaction vessel is not acceptable. Nor can the heat released by the MTG reaction be controlled within acceptable product gas exit temperature limits--i.e., 730.degree. F. to 800.degree. F.--by direct heat exchange methods, because the reaction MTG is too fast.
The need of an MTG catalyst for a sufficient, though limited, exposure to water coupled with its need for exposure to a sufficient, though limited, range of temperature to be sufficiently active, coupled to the fact that it catalyzes a fast and highly exothermic reaction, have entailed a number of processing complexities associated to the heat exchange requirements attendant to the proper operation of a methanol to gasoline process with an MTG catalyst.
In order to accommodate the temperature range exposure requirements of the MTG catalyst, the New Zealand facility effectuates product gas temperature control within the MTG catalyst vessels by recycle of the C.sub.2- and C.sub.3+ hydrocarbon by-product gases remaining after the reactor product gases are cooled down from their reaction temperature of about 800.degree. F. to about 100.degree. F. for removal of their gasoline grade hydrocarbon product content. In order to limit the temperature rise in the MTG reaction to less than about 100.degree. F., this recycle diluent gas is added in an amount compared to the methoxy feed gas (methanol, dimethyl ether, etc.) of about 12.4 moles diluent recycle/mole methoxy equivalent. For every 1 volume of methoxy equivalent with which it is combined to form the feedstock gas it is therefor necessary to heat about 12.4 volumes of the C.sub.2 and/or C.sub.3+ recycled diluent gas from its recovery temperature of about 100.degree. F. to the minimum temperature required by an MTG catalyst for sufficient activity, which is about 650.degree. F. and preferably about 700.degree. F. This imposes a tremendous heat exchange burden both in terms of reactor effluent gas cool down (800.degree. F. to 100.degree. F.) for gasoline product recovery and recycle gas (C.sub.2- /C.sub.3+) reheat (from 100.degree. F. to at least 650.degree. F., and preferably to 700.degree. F.). For each volume of hydrocarbon product produced and recovered about 71 volumes of recycle diluent gas must be heated up by about 600.degree. F. and later cooled down by about 700.degree. F. in adiabatic heat exchanger units by heat exchange first with a new portion of cool feed gas and then by cooling water or a refrigerant.
In the New Zealand operation the methoxy compound, generally in the form of an equilibrium mixture of methanol, dimethyl ether and water vapor, is combined with a recycle gas comprising C.sub.2- and C.sub.3+ hydrocarbons and the methoxy content is reacted in a single pass over an MTG catalyst to produce a product gas of about 800.degree. F. with a partial pressure of water (as steam) of 0 about 2 atmospheres absolute (ata) at a product gas pressure of about 22.3 ata. Hence, heat exchange operations to heat up feed gases and/or cool down product gas are carried out on gas streams at a pressure of about 22.3 ata.
The MTG reaction gases are cooled from their reaction temperature, T.sub.R, to their final cooldown temperature, T.sub.0, in two heat exchange operations. In the first operation the reaction gases are indirectly heat exchanged with feed gases through adiabatic heat exchangers to cause the feed gases to warm up from their unit available temperature of about T.sub.0 to their reactor inlet temperature T.sub.I by transferring thereto of a quantity of heat Q from the reaction gases which thus causes the reaction gases to cool down from their reaction temperature T.sub.R to their first cooldown temperature T.sub.x. Thereafter, the reaction gases are finally cooled from their first cooldown temperature T.sub.x to the liquid hydrocarbon recovery temperature, T.sub.0, either by direct or indirect heat exchange contact with chill water.
For purposes of discussing the heat exchange burden associated to operation of the New Zealand process, one may assume that the specific heat content of a feed gas about equals the specific heat content of the reaction product gas which results therefrom. Under this circumstance, when the reaction product gas gives up a quantity of heat Q to heat up a new portion of feed gas from T.sub.O to T.sub.I the reaction gas cools to a temperature T.sub.X =T.sub.R -(T.sub.I -T.sub.O). To finally cool the reaction gas to a temperature T.sub.O to recover its liquid hydrocarbon content the quantity of heat that must be removed from the reaction product gas is that additional quantity of heat resulting from the heat of reaction of the alkoxy compound from which the reaction gas was formed.
Generally, the adiabatic heat exchangers through which feed gas is heated up while reaction product gas is cooled down are on the order of 3-4 times as expensive as those heat exchange units by which the product gas is finally cooled down by heat exchange with chill water. Further, the sizing of each of these heat exchange units, which bears very significantly on their cost, is directly effected by the difference in temperature to which the feed gases are to be heated and the temperature from which the product gas is to be cooled (.DELTA.T) and the pressure (P) of the gases to be heated and cooled. The larger the differential temperature (.DELTA.T) and/or the higher the pressure the more efficient the heat exchange process and the smaller can be the size of the heat exchanger needed. The relative heat exchange surface area (A.sub.R) required to transfer this quantity of heat (Q) is generally given by the formula A.sub.R =Q (1/.DELTA.T)(1/P).sup.0.6.
Suggestions have appeared in the prior art which are directed to ameliorating this tremendous heat exchange burden. Mobil Oil Corporation's U.S. Pat. No. 4,035,430 to Dwyer et al. describes an MTG process wherein a number of spaced apart MTG catalyst beds are employed and hydrocarbon recycle gas may be injected between the successive beds to control the exotherm. Also described is the possibility of interbed injection of methanol or DME as a quench to maintain the temperature rise in each MTG catalyst bed to about 50.degree. F. with a total temperature rise over all beds to about 200.degree. F. This however, as Mobil's later U.S. Pat. No. 4,404,414 to Penick et al. notes, requires either that the conversion of the alkoxy feed in each MTG reactor be limited to a relatively low value (which is difficult to control at the high reaction rate in contact with the catalyst) or, alternatively, that an extraordinarily high recycle ratio must be used, neither being an attractive possibility.
Mobil's U.S. Pat. No. 4,404,414 to Penick et al. provides yet another suggestion. It describes the use of the reactor effluent gas of a preceding MTG reactor, for admixture with a fresh charge of alkoxy compound (DME and/or methanol) to be fed to a succeeding MTG reactor, as the medium for heating up the new alkoxy charge of combined gas mixture to the inlet temperature desired for reaction in a succeeding MTG reactor. According to Penick et al.'s suggestion the total quantity of temperature-control diluent gases used in the process is passed through the first MTG reactor and hence into the succeeding MTG reactor without the need for cooldown heat exchange of the diluent gas from the reaction gas temperature (about 800.degree. F.) to about 100.degree. F. for purposes of product recovery and thereafter reheat from about 100.degree. F. to an inlet reaction temperature (about 700.degree. F.) for recycle use with the new alkoxy charge. This reduces the overall effective volume of recycle gas as a ratio to the final product make required for proper temperature moderation in the MTG reactors. However, this heat exchange savings is purchased at the risk of advanced aging of the MTG catalyst in the succeeding MTG reactor that may result from the successive accumulation of water (steam) content that results from the buildup of water as a by-product of the MTG reaction that occurs within both the preceding and succeeding reactors. For example, for each mole of hydrocarbon produced, the hydrocarbon product from crude methanol having a water content of about 6 wt. % or from an equilibrium methanol-dimethyl ether mixture produced therefrom contains about 6.45 moles of water. Addition of a diluent gas to this methoxy compound in an amount sufficient to moderate the temperature rise in a first MTG reactor as desired, followed by addition to the first reactor product gas of an equivalent amount of new methoxy compound to form a feed gas stream for reaction in a second MTG reactor results in a product gas from the second reactor having nearly double a mole fraction of water as that of the product gas from the first MTG reactor.
Accordingly, in the process proposed by Penick et al. where both reactors are operated at about the same pressure the concentration of water (as steam) to which the MTG catalyst of the succeeding MTG reactors is subjected is substantially greater than that to which the MTG catalyst of the preceding MTG reactor is exposed. This, as Penick et al. notes, inevitably leads to the dilemma that if both reactors are operated at about the same temperature either the MTG catalyst in a preceding MTG reactor must be under-utilized by depriving it of exposure to a water content that optimizes its potential for activity or the MTG catalyst in a succeeding MTG reactor must be overexposed to water such that it ages prematurely. This dilemma of Penick et al.'s proposal necessitates a significant difference in the operational conditions of temperature or pressure between the preceding and succeeding MTG reactors, which, in order to maximize both the activity and operational life of the MTG catalyst in each, results in an undesirable change in hydrocarbon product's compositional distribution produced in each.
Yet another Mobil Oil Corporation U.S. Pat. No. 4,788,369 to Marsh et al. suggests a method for improving the economies of even the Penick et al. process that involves the separation of the C.sub.3+ by-product hydrocarbon gases from the C.sub.2- by-product hydrocarbon gases and thereafter use of the C.sub.3+ hydrocarbon gases as the recycle diluent gas required for temperature moderation. Because of the higher specific heat of the C.sub.3+ recycle, following the suggestion of Marsh et al. the volume of recycle diluent gas required for practice of the MTG process may be reduced. As Marsh et al. notes, in terms of compressor duties a C.sub.3+ recycle gas is more readily and economically processable for recycle use than is a C.sub.2- recycle gas. However, a smaller amount of the C.sub.3+ recycle gas volume necessitates use of a lower total pressure for processing if the steam vapor pressure is to be kept the same as before. That lower pressure has a negative effect on the heat exchange capability of the process. Thus, the only advantage remaining for the Marsh approach is the use of fewer moles of recycle gas which thus reduces compressor duties. The disadvantages of increased heat exchange surface requirements resulting from the lower pressure of operation reduces, and might even outweigh, the savings in the recycle compressor. It suffices to say that the Marsh et al. patent does not succeed in diminishing the heat exchange requirements of the process. Yet, even with use of a C.sub.3+ hydrocarbon by-product gas as a recycle to save on compressor duties as suggested by Marsh et al. the savings in capital and operational cost of the Marsh et al. process requirements are modest, at best.
The temperature rise/water concentration dilemma that hamstrings the MTG conversion process over an MTG catalyst as being an economically competitive process for gasoline production compared to that of gasoline refined from natural crude oil remains unsolved by any prior art suggestion. Given the abundance of natural gas reserves from which the production of gasoline grade products is technologically possible, a solution to the economical obstacle to its conversion to gasoline is a greatly desired goal.